Tesi Dottorato Scognamiglio

8/9/2019 Tesi Dottorato Scognamiglio

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UNIVERSITY OF STUDY OF NAPLES

FEDERICO II

Ph.D thesis in Chemical Engineering

(XXI Cycle)

HYDROGEN AS ENERGY CARRIER:DECENTRALIZED PRODUCTION BY

AUTOTHERMAL REFORMING OF METHANE

Relatore: Candidato:

Prof. Pier Luca Maffettone Ing. Diego Scognamiglio

Comitato scientifico

Prof. Silvestro Criscitelli

Prof. Gennaro Volpicelli

Ing. Lucia Russo

Ing. Marino Simeone


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Index

INTRODUCTION .................................................................................................................................................................. 7

HYDROGEN AS ENERGY CARRIER ............................................................................................................................... 7 PRODUCTION AND USES ................................................................................................................................................ 7

TRANSPORT AND STORAGE ........................................................................................................................................... 8

PRODUCTION PROCESS FROM METHANE ... ............................................................................................................ 11

STEAM REFORMING ....................................................................................................................................................... 11 THERMAL PARTIAL OXIDATION ................................................................................................................................. 14 CATALYTIC PARTIAL OXIDATION .............................................................................................................................. 17

AUTOTHERMAL REFORMING ...................................................................................................................................... 18 PROCESS CHOISE ............................................................................................................................................................ 21

EXSPERIMENTAL SETUP ................................................................................................................................................ 22

EXPERIMENTAL SYSTEM .............................................................................................................................................. 22

FEEDING SECTION ... ................................................................................................................................................. 22 REACTOR SECTION ... ................................................................................................................................................. 25 PRETREATING AND DISPOSAL SECTION ............................................................................................................. 26

ANALYSIS SECTION .................................................................................................................................................. 27 EXPERIMENTAL PROCEDURE ..................................................................................................................................... 28

REACTOR LOADING ... ............................................................................................................................................... 28 CATALYST PRETREATMENT ... ................................................................................................................................ 28 IGNITION ... ................................................................................................................................................................... 29

EXPERIMENTAL RUNS ... .......................................................................................................................................... 29

MATHEMATICAL MODEL OF FIXED BED REACTOR UNDER STEADY-STATE CONDITIONS ... ............... 30

INTRODUCTION ............................................................................................................................................................. 30 MATHEMATICAL MODEL .............................................................................................................................................. 31

GEOMETRY ... ............................................................................................................................................................... 31 MATHEMATICAL MODEL: MASS AND ENERGY BALANCE ... .......................................................................... 32 ESTABLISHMENT OF BOUNDARY CONDITIONS ................................................................................................. 34

ANALYSIS OF THE HYPOTHESES OF THE MODEL .................................................................................................. 35 STATE OF GASES ....................................................................................................................................................... 36

ANALYSIS OF THE PRESSURE DROP .................................................................................................................. 36 KINETICS IN HOMOGENEOUS PHASE ............................................................................................................. 37 HETEROGENEITY ... ................................................................................................................................................ 38 MASS AND HEAT DISPERSION IN GAS PHASE ... ............................................................................................. 40 DETERMINATION OF THE EXTERNAL HEAT TRANSFER COEFFICIENT ... ........................................ 44 PHYSICAL AND TRANSPORT PROPERTIES ...................................................................................................... 46

NUMERICAL METHOD ................................................................................................................................................... 46 DISCRETIZATION OF THE SYSTEM ... ..................................................................................................................... 47

RESULTS ON NICKEL BASIS CATALYST ... ................................................................................................................... 49

THERMODYNAMIC ANALYSIS ..................................................................................................................................... 49 KINETIC MODEL ............................................................................................................................................................. 52

MECHANISM AND KINETIC SCHEME .................................................................................................................... 52 EQUATIONS OF KINETIC SCHEME ......................................................................................................................... 54 KINETIC EVOLUTION OF THE PROCESSES .......................................................................................................... 54

EXPRESSION AND KINETIC PARAMETERS ... ....................................................................................................... 56 KINETIC DELAY: THEORETICAL AND EXPERIMENTAL ANALYSIS .................................................................... 59 EXPERIMENTAL RUNS AND COMPARISON WITH THE SIMULATED DATA ........................................................ 63


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RUNS VARYING THE FLOW RATE ......................................................................................................................... 64 RUNS VARYING THE FEED COMPOSITION ... ....................................................................................................... 68

OBSERVATIONS ON MATHEMATICAL MODEL ........................................................................................................ 73 TRANSPORT PHENOMENA ANALYSIS .................................................................................................................. 74 IMPORTANCE OF IR ANALYSIS: HETEROGENEITY ... ........................................................................................ 78

EFFECT OF PRESSURE .............................................................................................................................................. 81

EFFECT OF PREHEATING TEMPERATURE ... ....................................................................................................... 83 LIMITS OF MATHEMATICAL MODEL ... .................................................................................................................. 84 CONCLUSIONS ON NICKEL CATALYST ...................................................................................................................... 85

RESULTS ON RHODIUM CATALYST ... ......................................................................................................................... 87

KINETIC MECHANISM: DIRECT OR INDIRCT? .......................................................................................................... 88 ASSUMPTION: INDIRECT KINETIC MECHANISM ... .............................................................................................. 90 EQUATIONS AND PARAMETERS OF KINETIC REACTIONS ... ........................................................................... 90 KINETIC EVOLUTION OF CHEMICAL PROCESSES ... ........................................................................................... 93

MATHEMATICAL MODEL AND KINETIC HYPOTHESIS VALIDATION ................................................................ 94 EXPERIMENTAL RUNS AND COMPARISON WITH SIMULATED DATA ................................................................ 97

RUNS AT DIFFERENT VALUE OF AIR TO CARBON FEEDING RATIO ............................................................ 97 AUTOTHERMAL REFORMING RUNS .... ................................................................................................................ 100

THEORETICAL EFFECT OF OPERATIVE FEED CONDITIONS ON KINETIC MECHANISM .... ........................ 104 RUN1: CPO WITH A HIGH PREHEATING TEMPERATURE AND A LOW FEED RATIO Air to Methane .... ..... 106 RUN 2,3: ATR WITH HIGH PREHEATING TEMPERATURE AND LOW FEED RATIO Air to Methane. .... ....... 108

CONCLUSIONS ON RHODIUM CATALYST .... ........................................................................................................... 111 COMPARISON BETWEEN NICKEL AND RHODIUM CATALYST UNDER CPO AND ATR PROCESSES OFMETHANE .... ...................................................................................................................................................................... 112

DISTRIBUITION OF KINETIC ZONE .... ..................................................................................................................... 112 THERMAL PROFILES: HOT SPOT AND MAXIMUM TEMPERATURES .... .............................................................. 116 PERFORMANCE: CONVERSION AND YIELD .... ....................................................................................................... 120 CONCLUSION ON COMPARISON BETWEEN NICKEL AND RHODIUM CATALYST .... ..................................... 123

DEVELOPMENT .... ........................................................................................................................................................... 125

EFFECT OF CATALYST TYPE: NUMERICAL INVESTIGATION .... .......................................................................... 125 NON CONVENTIONAL REACTORS .... ....................................................................................................................... 127

MEMBRANE REACTORS .... ...................................................................................................................................... 127 CATALYTIC WALLS REACTORS .... ........................................................................................................................ 129 FORCED REACTORS .... ............................................................................................................................................. 131

NOMENCLATURE .... ........................................................................................................................................................ 138

SYMBOLS.... ..................................................................................................................................................................... 138 GREEK SYMBOLS.... ...................................................................................................................................................... 138 SUBSCRIPT .... .................................................................................................................................................................. 138

REFERENCES .... ................................................................................................................................................................ 139

FIGURES INDEX .... ........................................................................................................................................................... 144


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INTRODUCTION

HYDROGEN AS ENERGY CARRIER

Hydrogen is the most abundant element in the universe (as recently demonstrated by theHubble Space Telescope, that has revealed great quantity of this substance), astrophysicalsaid, although it is not found in a pure state (hydrogen deposits do not exist). It wasdiscovered in 1766 by Cavendish who named it inflammable air," then Lavoisier gave theelement the name of hydrogen (from the Greek hydro meaning water and genes meaningcreator).Hydrogen has been employed as fuel for domestic use for more than a century: the so-called city gas, in fact, consists mainly of hydrogen and carbon monoxide (about 50 %).

The diffusion of this gas has been limited due both to toxicity of carbon monoxide and tothe increased use of methane as fuel.Nowadays, the name of hydrogen, which is put beside to that of fuel cell, is the key wordof energy as well as the fulcrum of financial support and scientific interests. The scientistsdefine hydrogen the fuel for excellence, the cleaner method to store, transport and utilizethe energy, an excellent energy carrier; to this aim it can be remind that the hydrogencombustion produce only steam and energy:

H2 + O2 = H2O + energy

Hydrogen is an ideal complement to the electricity, in fact, both are premium quality energycarriers, do not contain carbon and generate little or no polluting emissions at the point ofuse. The electricity, however, presents some disadvantages concerning the storage whereasthe hydrogen is more satisfactory and its versatility is demonstrated from the possibility toobtain it from a variety of primary source that can be fossil or not. So, a hydrogen supplyinfrastructure could be developed, according to available reserves in every region. Thiscould provide concrete benefits to energy security of supply and could ease the transitiontowards a sustainable energy system.

PRODUCTION AND USES

Hydrogen can be generated both from renewable source and from hydrocarbon fuels. It would be better to use renewable source; in this case, the developing countries would beadvantaged due to their abundant resources available, also if they dont have the necessaryinfrastructure. Instead, in industrialized countries where the availability of infrastructure is


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higher (e.g. natural gas distribution network), the non-renewable resources probably will beutilized at least in the beginning.Hydrogen industrial production is based on well-known processes that are utilized until acentury. Current world hydrogen production is approximately 45 million tons/year, 90 % ofthese was generated from fossil fuels (oil, carbon and natural gas). The largest use ofhydrogen in the world is to manufacture nitrogen fertilizers which follow the use in thepetrochemical processes. World hydrogen consumption is due to 50% for ammoniaproduction, 37% for oil refinery, 8% for methanol production, 4% for general uses and 1%for aerospace activities. Also if one admit that the hydrogen production from renewablesources (ruling is fixed from Kyoto protocol, and obviously, out the fossil and/or nuclearsources) could not be economically competitive in the immediate future, the hydrogenproduction in controlled conditions (even coupling the process to another for CO2sequestration) could allow a reduction of emissions of polluting substances, especially in thehighly urbanized zone characterized by intense truck traffic. This could contribute to the

reaching of emission limits stabilised from Kyoto protocol and, obviously, rendering morebreathable the air in our city. However, it is important to highlight that hydrogen is anenergy carrier and not an alternative fuel source as wrongly affirmed. Therefore, it isevident that an hydrogen-based economy would need to determine the optimal utilizationof hydrogen (from economical and technical point of view) and to resolve the problemsrelative both its production and global environmental impacts.

TRANSPORT AND STORAGE After the production, hydrogen has to transport or store properly in various aggregation states: solid, liquid or gas. This delicate operation represent one of the most importantproblem that have to be resolved because the transition to a hydrogen based economybecoming possible. In fact, also if is true that hydrogen is, from gravimetric point of view,the substance that allow the best ratio between stored energy and weight (about 2.5 timehigher than methane), it is necessary consider that it have a low energy density (about 3.2time lower than methane), an high volatility and a dew point -250C. Not negligible arealso the safety issue due to its extreme reactivity and to the explosion danger. The USDepartment of Energy (DOE) established that hydrogen will be considered an acceptablefuel when the following target will be fulfilled:

Energy density of 6.5% wt. of total drum weight in hydrogen; Volumetric density of 62 Kg/m3.

These targets are based on hydrogen fuel cell vehicles with autonomy of 560 Km. Actuallyfor hydrogen storage there are three fundamental options:


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Compressed gas in pressurized cylinders; Liquid in vessel at low temperature and pressure; In solid solution with other compounds, generally as simple or composite metal

hydrides.

Hydrogen is commonly stored in high pressure cylinders (200-350 bar). The heavyindustrial cylinders can be usefully replaced with dual walls ones made of aluminium. Understudy there are cylinders made of carbon fibres that allow pressure of about 700 bar also ifthese devices are still very heavy for the utilization as on-boards tanks. The disadvantage ofthis system is that hydrogen form hydrides with metal resulting from the internal wall of thecylinder weakening its structure.For the liquid storage considering that is not necessary that tanks both undergo highpressure and have an high weight (generally the thermal insulation dont require materials

with an high specific weight) it can be easily reach a weight ratio that satisfy the DOE target

(6.5%), as long as the system is adequately insulated in double wall tanks provided of cavity wall insulation because the hydrogen condensation occur at -253C; moreover the lowchemical reactivity of hydrogen at low temperature eliminate the problem of metal fragility.

This technology has been widely developed for aerospace missions (NASA, ESA) andhydrogen, in liquid phase, has been considered very interesting for high distance transportas fuel for planes and rockets.It is necessary to take into account the cost relative to hydrogen condensation (the energeticcosts are about 30-40% of the liquid intrinsic energy) to thermal insulation and the safetyissue.

Another transportation method is based on the capability of hydrogen to form hydrides;this reaction often occur spontaneously still at room temperature and proceed to right orleft depending on the gaseous hydrogen pressure. The advantages from the use of hydridesare the greater safety and the extremely high density that allow to reach; the disadvantage isrepresented from the high temperatures (150-300C) required for the hydrogen extractionfrom metal hydrides and, obviously, from the relative cost.In the future, in this research field, the work will be addressed toward the production oftanks made of composite materials for compressed hydrogen and complex metal hydrides,

whereas a new technique under experimentation is the hydrogen adsorption in carbonnanostructures instead of the hydration of sodium boron-hydrate and the encapsulation inglass microstructures.For the utilization as energy carrier, the hydrogen has to be easily transportable in safetyconditions minimizing gas losses. A possible transport system is represented from hydrogenpipeline equivalent to methane pipeline but with an improving of used materials. Analternative could be build decentralized hydrogen production plants starting from methane,so leaving unchanged the methane pipeline.


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The aim of this PhD thesis will be the mathematical modeling, followed by an experimentalcampaign designed at evaluating the performance of an autothermal reactor for thehydrogen production from methane. The catalysts used are of commercial pellets type andbecause the think it substantially to a reactor for automotive and/or residential purposes, itmust be as compact as possible.

The attention will be focused on thermal profile which is established within the catalyticbed so that it can evaluate the thermal stress to which the catalyst undergoes (the mainproblem that inhibits the development of autothermal processes on catalyst is the thermaldeactivation of its due to high temperatures reached). In this regard will be, experimentally,determinate the thermal profile with IR technology so as to make possible a reliable

validation of the mathematical model and then simulate the behavior in a situationprohibitive and/or to evaluate off-line the effect of process variables.

The results will clearly also used for a comparison between the two commercial catalystsselected.

In this regard, into the following chapters will introduce the processes of production ofhydrogen from methane starting from the already noted industrial technologies for theproduction on large scale, highlighting the limit for their development from the perspectiveof decentralized production until the processes currently being experimentation. It will beshown the experimental apparatus and procedures used in the laboratory. It will thenillustrate the experimental and numerical results obtained on the two catalytic systems withits comparison, and finally there will be a small section devoted to ideas for possible futuredevelopments as well as an overview of literature on non-conventional reactors that it canpossibly utilized for such a process.


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PRODUCTION PROCESS FROM METHANE

STEAM REFORMING

The process generates hydrogen from steam and methane according the followingreactions:

1) CH4 + H 2O CO + 3H 2 H (298) = 206000 KJ/Kmol (reforming of CH 4 )

2) CO + H 2O CO2 + H 2 H (298) = -41000 KJ/Kmol (CO shift)

3) CH4 C + 2H 2 H (298)= 76000 KJ/Kmol (cracking of CH 4 )

The overall process is endothermic and catalytic (generally, the catalysts used is nickelsupported on alumina) and the reaction proceed with an increase of total number ofmolecules; so a thermodynamic consideration suggest that hydrogen production ispromoted at high temperatures (T), low pressures (P) and high ratio steam/methane (S/C).

Figure 1 : Sketch of reforming reactor .

The endothermicity of the process require the heating of the system at high temperaturesby using a shell-and-tube reactor. The methane and the steam are fed to tubes that are


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dipped into a chamber where there are burners located in co-current with the feedingcurrent. It can be noted that the steam reforming reaction is favoured by high temperature

whereas the CO shift ones is favoured by low temperatures; this implies that at the out ofthe reforming reactor the current have to be send to a CO shift reactor that work at lowertemperatures so that it can be obtained a rich hydrogen current.Generally, it have to be determined the three parameters T, P, S/C that maximize thehydrogen production in an acceptable temperature field without the coke formation.

Acceptable temperatures means that it are technologically compatible with tube materials(metal materials not ceramic ones to allow the thermal exchange) and with catalyst thatrequire temperatures lower than that of sintering, which determine lowering specific surface

with a consequent lowering of catalyst activity. Thermodynamic of steam reforming has been studied by some researchers with the aim toidentify the best values of three parameter above mentioned.

Y.S. Seo et al. [1], by means of Aspen Plus software, found three values of parameters that

optimize the hydrogen yield, minimizing the CO quantity, working at a temperature that notallow both the catalyst deactivation and the formation of coke.In Figure 2 . is shown the effect of temperature on molar fraction of various products alsoindicating what is the zone under conditions for C formation. Being limited to a T of about800 C (for technologic reason), with stoichiometric feed, it is not possible obtain aconversion degree higher than 90%. In Figure 3 is shown the variability of molar fluxes of H2, CO and C with the ratio S/C; itcan be noted that rising this ratio increase as C decrease until to a total disappearance for a

value of 1.4, moreover the increase of S/C determine an increase of H 2 flux and a decreaseof CO one.

In2

Figure 4 is shown the influence of temperature and o f the ratio S/C on conversiondegree; it can be noted that if it is wanted a conversion degree of 99% at 800C, it isnecessary reach a feed ratio S/C higher than 1.9. However, it is necessary remind at anincrease of S/C is expensive.Finally, the necessity of heating and the long time for ignition and turndown of the processobstruct the realization of this operation on moving vehicles or on small scale decentralizedproduction unity. Nevertheless, steam reforming allows obtaining the best yields insynthesis gas with respect to other processes, being the heating source external to thereactor.


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Figure 2 : Influence of reactor temperature on composition and conversion degree at the pressure of 1 bar and r1.

Figure 3 : Influence of the ratio S/C on the equilibrium composition; reactor at pressure of 1 bar and tempera600C (-----), 700C ( ) and 800C ().


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Figure 4 : Influence of the ratio S/C on the conversion degree; reactor at a pressure of 1 bar and temperatu

600C (-----), 700C ( ) and 800C ().

THERMAL PARTIAL OXIDATION

In steam reforming the necessity of an external heating obstruct the reproducibility in thedecentralized production plants. To resolve this issue it could be heating, during theendothermic reaction, by means the develop of exothermic reactions in the same reactor.

This just happened in the partial oxidation: the process is carried out without catalyst,feeding methane and oxygen in a reactor made for work in adiabatic conditions. In the firstpart of the reactor (which is tubular ed, obviously, is without the burner) take place theoxidative processes that generate heat and steam for reforming reactions that are carriedout in the second part of the reactor until the thermodynamic equilibrium is reached. Thefundamental parameter in partial oxidation is the ratio O2/CH4; in presence of methaneand oxygen, the oxidative processes that could be take place are:

1) CH4 +0.5O 2 CO + 2H 2 H (298)= -36000 KJ/Kmol

2) CH4 +2O 2 CO 2 + 2H 2 O H (298) = -802000 KJ/Kmol


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A value of the ratio O2/CH4 of about 0.5 promote the formation of products from partialoxidation with respect to total combustion ones but prevent the reaching of hightemperatures in autothermal mode. So, the unreacted methane instead to react with steamaccording to reforming reactions generating H2 and CO, in part pyrolize forming carboncoke and the other part remain unconverted with a low yield in synthesis gas. For to reach,in autothermal mode, the high temperatures required for the maximization of yield insynthesis gas, it is necessary that system works at ratio O2/CH 4 higher than 0.5. In this waythe total combustion reaction also take place lowering the selectivity but generating heatthat allow the reaching high temperatures in autothermal mode.So, in the partial oxidation the endothermicity problem is solved but this means a loweringof selectivity, higher CO content and high temperatures due to the scarce overlappingbetween oxidation and reforming zones.Seo et al. [1] also studied the thermodynamic of partial oxidation. The stoichiometriccoefficient of oxygen was varied from 0.0 to 1.2 (in Seo et al.s work the air ratio was defined

as the half of that of oxygen).In Figure 5 are shown the composition of various equilibrium products at varying the airratio at preheating T of 200C and a P of 1 atm. It can be noted that for ratio O2/CH4higher than 0.6 ( air ratio higher than 0.3) there is not carbon coke formation, but it can beseen also a reduction of H2 and CO yield, with a reduction of H2 formation higher than aCO.In Figure 6 is reported the monotonic increasing trend of the degree of methaneconversion, x, and of adiabatic temperature with the air ratio at preheating T of 200C and aP of 1 atm; from the same figure it can be also noted that hydrogen yield show a maximumfor values of air ratio close to 0.3.

It is important to remember that the reactor of partial oxidation is more compact to respectto that one utilized for the steam reforming, in which there is an external heating.From energy efficiency point of view, the yield of partial oxidation is about of 50%,

whereas in the steam reforming is about of 65-70% and this depend both upon the highprocess temperatures and upon the problem of heat storage inside the reactor [2].

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Figure 5 : Equilibrium composition at varying the air ratio; the preheating temperature of methane and air is 20the pressure is 1 atm.

Figure 6 : Adiabatic temperature, degree of methane conversion and hydrogen yield, at varying the air ratio; thetemperature of methane and air is 200C and the pressure is 1 atm.


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CATALYTIC PARTIAL OXIDATION

In the last decade, many researchers have devoted their attention to catalytic partialoxidation. This technique allows working at lower temperatures with respect to thermal

partial oxidation, utilizing air instead of oxygen and with very small reactor due to greaterreaction rates. In fact, while the steam reforming contact time is of the order of seconds,those for catalytic partial oxidation are of the order of milliseconds.Unlike to steam reforming, for this process the catalysis is continuously improved andtoday does not still exist a standard reference catalyst. In literature, there are many worksabout different catalysts focused to the searching of catalytic solution that allows theoverlapping between reforming and oxidation reactions. This overlapping is essential toobtain, inside to catalytic bed, a thermal profile as smooth as possible, without peaks thatcould submit the catalyst a strong thermal shock.In particular, for this process, it is not very clear what the reaction mechanism is. There are

two great schools of thought that born to the direct and indirect reaction mechanismproposed by Prette et al. [3] e Schimidt et al. [4].

The indirect mechanism hypothesize that the primary products are CO 2 and H 2O producedfrom the total combustion of a part of methane that react according this reaction until alloxygen is consumed, and that the secondary products are CO and H 2 produced byreforming of un-reacted methane with CO 2 and H 2O previously formed. Conversely, themechanism proposed from Schimidt hypothesizes the direct partial oxidation of methane toCO and H2 by means of methane pyrolysis to C and H2 and successive oxidation of C toCO.Many researchers think that reaction mechanism depend upon the types of catalyst used; inparticular from both metal and support types.

Yan et al, on 2004 [5] studied the reaction mechanism of partial oxidation of methane ontwo catalyst that have the same support but different active principles: Rh/SiO 2 andRu/SiO 2. They were shown that on Ru/SiO 2 the partial oxidation took place according tomechanism proposed by Prette, whereas on Rh/SiO 2 according to mechanism proposed bySchmidt. The Authors justified these results on the base of the different bond strength ofM O.

The strong bond Ru O prevents the reduction of metal during the partial oxidationprocess favouring the formation of the products derived from the total combustion. The

bond Rh O instead is very weaker and it facilitates the reduction of Rh promoting thedissociation of methane and the successive formation of products from partial oxidation. The Authors of this work studied also the reaction mechanism of partial oxidation on twocatalysts characterized by two different supports but having the same active principle:Rh/SiO 2 and Rh/Al 2O3. They were shown that the partial oxidation on Rh/Al 2O 3 tookplace toward an indirect mechanism whereas on Rh/SiO 2 toward a direct mechanismjustifying this on the base of different interaction between active principle and support: in


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fact, a strong interaction inhibits the metal reduction favouring the formation of productsfrom total oxidation.

Veser et al [6] claim that the reaction mechanism of partial oxidation is always direct type,independently of catalyst types used. They think that the reason for which, at beginning,great quantities of CO 2 and H 2O formed is due to different values of adsorption coefficientof oxygen and methane and not to the indirect kinetics mechanism. Since the adsorptioncoefficient of oxygen is higher than of methane, at beginning, the catalytic surface is coverpreferably from oxygen, promoting the total combustion reaction; during this phase moreoxygen is consumed with respect to methane and this involved a lowering of ratio CH 4/O 2 gas phase. When the partial pressure of O 2 is as low as compensate the different values ofadsorption coefficients, on the catalytic surface there will be combustion conditions that arerich of fuel favouring the formation of products from partial oxidation.However, beyond of kinetic mechanism, the developing catalytic of partial oxidation in theindustrial technology field is actually inhibit from the presence of hot-spot (temperature

peaks) inside to the catalytic bed that shorten the catalyst life as well as the cost of catalyst.

AUTOTHERMAL REFORMING

Autothermal reforming is a combination between steam reforming and catalytic partialoxidation. In fact, the feed is constituted from water, oxygen and methane with feedingratio such a process is autothermal and that there is not the formation of carbon coke. Like

so the partial oxidation, also in this process there is a heating inside the reactor: in the firstpart the exothermic oxidation reaction take place and furnish heat for the endothermicreaction of reforming.

The addition of steam allows obtaining lower temperatures and higher selectivity tosynthesis gas with respect to that obtained in partial oxidation. Moreover, this additionlimit the formation of carbon coke and the presence of hot-spot in the catalytic bed becausereforming reaction are anticipated and so overlap better the oxidation reaction. Clearly,addition of water extend the contact times with catalytic bed (about one tenth of second)and so it is not possible utilized compact reactor as well as that utilized in the oxidationprocess but however extremely compact with respect to reforming ones due to the insideheating.Seo et al. [1] (2002) studied the thermodynamic of autothermal reforming process ofmethane at atmospheric pressure and with preheating temperature of 400C.In Figure 7 it can be noted that conversion degree deep increase with increasing the airratio, until 1 when the air ratio is 0.3; the temperature increase with air ratio also for valueshigher than 0.3, because also if the conversion of methane reach 1 , H2 and CO can be stilloxidized. Increasing the ratio S/C, the x and T decrease because the steam reforming


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processes is more important and imply an higher heat consuming that induce a lowering oftemperature and so also of the conversion degree.

Figure 7 : Effect of air ratio and S/C on adiabatic temperature and conversion degree; reactor work with a pressu

and a preheating temperatureof 400C.

In Figure 8 is shown the formation of carbon coke depending on S/C and on air ratio. Itcan be noted that increasing S/C the carbon coke formation zone becomes smaller.From Figure 9 it can be noted instead, that molar flux of hydrogen and CO increasing withS/C, whereas their trend with air ratio is non-monotonous with a maximum around 0.25.


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Figure 8 : Effect of air ratio and S/C on carbon coke formation zone; reactor work with a pressure of 1 bar a preheating temperature of 400C.

Figure 9 : Effect of air ratio and S/C on the molar flux of hydrogen and carbon monoxide.


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In the autothermal reforming the oxidation reaction can take place, in absence of catalyst,utilizing the heat and steam produced for catalytic reforming reactions (this process work atindustrial scale). In this process, it also possible that oxidation reaction take place on acatalyst that can be or not the same utilized the reforming process. The choice of utilize twodifferent catalysts (Pt for oxidation and Ni for the reforming reaction) ameliorate thethermal management of process favouring the overlapping between the oxidative andreforming zone. The overlapping is further enhanced when the two active principles areimpregnated on the same support.In this field, the catalysis is continuously improved with research direct to develop catalyst

with high performance, but at industrial scale the autothermal reforming process run withnickel-based catalyst.Moreover when methane, oxygen and water are put in contact with a catalyst differentreaction can be occur in the reactor; in spite of the complexity of this system, few works aredevoted to the kinetic study on autothermal reforming ( [7], [8] ) and in these study it is

supposed that the autothermal reforming take place according to the indirect mechanism.

PROCESS CHOISE

Once done overview of the possible processes used to hydrogen production from methane,it seems clear that, with the expressed purpose of this thesis (detailed study of the catalyticreactor to be used in optical of decentralized hydrogen production and/or small-scale

processes), the candidate to the development of such a reactor are clearly those autothermic(CPO and ATR). Infact they are developed completely on a catalytic and the energy neededto sustain the reactions reformer, is generate the same reactor.For these reasons, the two commercial catalysts taken into consideration, will be analyzedfor test of catalytic partial oxidation and steam reforming.


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EXSPERIMENTAL SETUP

In the following sections will be described the experimental apparatus utilized in theautothermal reforming runs. Moreover, it will be illustrated the program of experimental

runs that will be described and discussed in the following chapters and utilized for the validation of the proposed mathematical model.

EXPERIMENTAL SYSTEM

The experimental system utilized is described Figure 10 It is constitutes from:

Feeding section; Reactor section; Pre-treatment and gas disposal section; Analysis section.

FEEDING SECTION

Feeding section is constituted from three cylinders at high purity provided with pressureregulators that allow a delivery pressure of 0-6 Bar, for the feeding of oxygen, methane andnitrogen. The regulation of gaseous reactants is done by means a Mass Flow Controller(MFC) purchased from Brooks with the following characteristics:

Model MFC 5850;

Gas: Methane; nitrogen; oxygen; Flow range: 0-5 l/min; Maximum temperature: 65C; Maximum pressure: 100bar; Serial communication :RS232; Feeding: 24V DC; Accuracy: 0.7% vl; 0.2%fs.


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The steam are produced in situ feeding water from a vessel opportunely pressurized withnitrogen; the liquid water flow from vessel is controlled by a Mass Flow Controllerpurchased from Bronchorst, with the following characteristics:

Model MFC L23V02; Flow range: 1-300 gr/h; Maximum temperature: 65C; Maximum pressure: 35bar; Serial communication: RS232; Feeding: 24V DC; Accuracy: 0.7% vl.

The vaporization of water and gas preheating is done by means of an evaporatordownstream of MFC. The evaporator is constituted of two 316 stainless steel plates, in oneof this plate occur the preheating of gas CH4, O2, N2 whereas in the other occur the

vaporization of water. In the plates there are due passing holes that allow the feeding flowand six blind holes to lodging six cartridge heaters (100W, 220V) that heating the plate. Perto contain heating loss, the evaporator is insulated with rock wool.

The gas stream constituted of CH4, O2 e N2 pass through the first plate that function aspreheater and then is joined to water passing through the second plate that functions as

full-blown evaporator. The feeding of water inside the evaporator is done by means a diptube (length 6 cm, inner diameter 1 mm) that is the inner tube of a concentric tube heatexchanger; instead in the annulus pass the preheating gas stream that works also as carriergas for the steam produced in the evaporator channels. The mixing of the gas stream withsteam allows limiting the pressure fluctuation deriving from local variation of temperature.Inside the evaporators, little Pyrex glass spheres are introduced to also ameliorate themixing between steam and carrier gas.

The temperature of plates, and so of the outlet gas, is controlled by a PID equipped with aninlet for thermocouples K type and an outlet power relay connected to the feeding ofcartridge heaters.

To prevent the condensation of water in the steel tubes that are connected with theevaporator, strip heaters are utilized to allow that the feeding stream enter in the reactor attemperatures higher than 100C.


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Figure 10: Scheme of experimental system.

N2

O 2

ABB

GC

H 2O

E

R

T

F 2

F 3

F 4

F F

A2

F 1

C

TC

H2O, IN R1

V3

V4

H2O, OUT

N2

V

IR

R

Cappa

CH 4

F1: MFC per H2OF2: MFC per N2 F3: MFC per O 2 F4: MFC per CH 4 E: Evaporator

A1,A2: AsameterIR: Thermocam

FP: FurnaceR: Reactor

V 1: Valve 4 way V 2: Valve 3 wayC: Condenser

T: Chemical trap TC: Catalytic torch V: Vessel for water


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REACTOR SECTION

The reactor sectionis constituted from the reactor and the furnace that contain it. To studythe autothermal reforming process is utilized a tubular reactor with a gas inlet and outlet,equipped of a lodging for a thermocouple.

The choice of material for the reactor is done taking into account the high temperaturesnecessary for reaching the wanted yield in synthesis gas and the nature of gaseous species.

The necessity to operate at temperatures higher than 800C reduce the possibility of choiceat only quartz, incoloy and steel; the last is discarded because the formation of hydrides dueto the contemporary presence of hydrogen inside the reactor and of high temperatures.Instead, the incoloy, that is an alloy of nickel and chromium often utilized in the operationin which there are high temperatures and hydrogen in gas stream, could be utilized for theconstruction of our reactor but the presence of nickel in this alloy can be foul the measurescarried out catalyzing the reforming reaction of methane.

For these reasons, it has been utilized a reactor made of quartz that resist at temperatureshigher than 1200C and that is inert toward the different compounds of gas stream. Thesole difficulty in the utilization of quartz is due to the low resistance at strong thermaloscillation; this render necessary a specific procedures for the start up and the stand-by ofthe reactor avoiding to expose it to strong thermal shock that could lead to its breaking.

The tubular reactor is having an outer diameter of 25 mm (thickness of about 2 mm) andfrom a total length of 60 cm; gas inlet and outlet having an outer diameter of 6mm. Only alittle part of the reactor (e.g. 14 mm, corresponding to 5 gr of Nickel catalyst in pellet ofabout 1 mm) is taken from the catalytic bed, the remaining part is filled with an inertmaterial of the same type of catalytic support (Al 2O 3 with size of about 1 mm). Thismaterial is utilized both for preheating of reactants at the inlet of the catalytic bed and toreduce the issue of possible flame propagation from mixture of methane and oxygen whichconcentrations fall in the flammability limits; moreover it can function as thermal barrierand capture the radiation emitted from catalytic bed so increasing the thermal efficiency ofthe process.

The insulation of the reactor is realized by a cylindrical block made of ceramic material witha length of 15 cm, inner diameter of 2.5 cm a width of 2 cm.

As above mentioned, the reactor is placed inside a vertical furnace, having a length ofheated zone of 40 cm and a reaching maximum temperature of 1200C. The furnace isequipped with a PID temperature controller, connected to PC by means of RS232, thatallow to reach the final temperature in a set mode (with appropriate heating ramps) byutilizing an algorithm realized in Labview environment, so avoiding strong thermal shock tothe quartz.Pressure transducers with a measuring range of 0-6 bar and analogic output measure thepressure upstream and downstream of the reactor.

An overpressure valve has been put upstream to the quartz reactor to avoid the broken inthe case of anomalous pressure increasing (release pressure of about 2.5 bar).


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The experiments are carried out on two different commercial catalytic systems, purchasedfrom Engelhard, containing nickel e rhodium respectively and both supported on alumina.Chemical-physical properties of these catalysts are reported in Table 1 .

The nickel catalyst pellets have the shape of cylinder with a diameter of 16 mm in whichthere are four axial holes. To minimize the fluidodynamic deviation in the catalytic bed, thecylindrical pellets are fragmented to the particles with diameter in the range 11.18 mm andthen are sieved and utilized as catalyst. In the experimental runs, it has been utilized 5 g ofcatalyst corresponding to a length of 14 mm of the catalytic bed. The rhodium catalystpellets have a diameter in the range 11.18 mm. Two quartz disks, with a diameter equal tothe inner part of the reactor, divide the catalytic zone from the inert one.

Table 1: Chemical-physical properties of catalysts.

PRETREATING AND DISPOSAL SECTION

The pretreating and gas disposal section is constituted from a water and particulate eliminationsystem from the gas stream outlet of catalytic reactor and from a unburned disposal system.

A part of the gas stream is sent directly to gas disposal system where undergo a treatmentthat is different from that reserved for the other part sent firstly to analysis section and then

to disposal unity. A chemical trap, CaCl2, carries out the elimination of water in the gas stream sent toanalysis section that is highly selective towards water molecules. The solid particulate iseliminated from gas stream by means a filter of 45 micron located downstream to thechemical trap.Instead, the part of gas sent directly to gas disposal section pass inside a condenser toreduce water concentration until to a value corresponding to the vapour pressure at thetemperature of cooling fluid (water). The presence of condenser is due to the possibility

NICKEL RHODIUM

Composition (wt%)NiO 16.5 Rh 0.4CaO 6.0 Al2O 3 Balance

Si 0.1 Al2O 3 Balance

Physical propertiesBulk density (kg/litre) 0.95 Bulk density (kg/litre) 0.8

Crush strength (N/cm) 400 Total surface area (m^2/g) 10Size of the spheres (mm) 11.18 Size of the spheres (mm) 11.18


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that the water contained in gas outlet of the reactor could condensate crossing coolpipeline, blocking it.

To avoid that part of formed CO2 is absorbed in the condensate (this can occur in physicaldehydration process), so fouling the experimental measurements that involve the CO2, thepart of gas stream that have to be sent to analyzer do not crossing the condenser but it issent directly to the chemical trap.Finally, the dehydrated gas stream is send to a catalytic torch where take place the oxidationreaction of all unburned contained in the gas stream. This torch is represented by a vessel

with a diameter of 5 cm and length of 15 cm, equipped with two inlets (one for air and onefor the stream outlet of the reactor), and an exit. Inside the torch it has been introduced amonolithic catalyst, purchased from Infragas, in which the active principle is platinum (3%

wt) whereas the support is made of cordierite (200 cell per square inch). The monolith isintroduced as to not leave empty space for gas passage.

ANALYSIS SECTION

The analysis section is constituted from two parts, one for the analysis of gas composition,both in inlet and outlet from catalytic reactor, and one for the acquisition of thermal profileof solid phase inside the reactor.For the analysis of gas composition is utilized a continuous analyzer, purchased from ABB,that is composed of three modulus:

- The first, Caldos, is an analyzer based on the measurement of thermal conductivity tomonitoring the hydrogen concentration;- The second, URAS, is an infrared analyzer to monitoring methane, carbon dioxide and carbon monoxide;- The third, MAGNOS, provided of four measure fields, to monitoring oxygenconcentration.

These units are provided of software for the interference corrections on the measurementof hydrogen due to presence of the other substances.

The thermal profile inside the catalytic reactor is measured by means of a infrared (IR)thermocam (Phoenix, Flir System) that allow to acquire the radiation emitted in a

wavelength range of 2-5 m, radiation emitted from solid inside the reactor and not fromexternal quartz wall, with a resolution of 320*256 pixels. This thermal profile is afundamental measurement both for the model validation and for the study of catalystbehaviour under typical feeding of autothermal reforming.

All the analogical signals, which those coming form thermocouples, pressure transducers,etc., are sent to a device from National Instruments (model NI-PCI-6229) for dataacquisition and digitalization by means an appropriate Labview algorithm.


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EXPERIMENTAL PROCEDURE

The experimental procedure is divided in following phases:

Reactor loading; Catalyst pretreatment; Ignition; Experimental runs.

REACTOR LOADING

The reactor has been loading with 5 g of nickel-catalyst supported on alumina, reduced inpellet of 1-1.18 mm to avoid the formation of preferential flux lines. The catalyst is put oncentre of the quartz tubular reactor, occupying a length of about 14-20 mm, between tworegions of inert material (about 12 cm of alumina of the same size of catalyst) to guaranteethe reactant mixing and to collect the thermal radiation emitted from catalytic bed. The loadreactor has been insulated by means a ceramic block and inserted inside the furnace that isindispensable for pretreatment phase, for start up and for preheating of gas stream.

CATALYST PRETREATMENT

After the loading and the placement of reactor there is the reduction in situ of the catalyst,that is indispensable because it has been oxidized by atmospheric oxygen during the loadingphase.

The reduction take place treating catalyst with a stream of H2 in N2 (30% in H2), attemperature of 600C, for about one hour. This temperature is reached, under gas flux,starting from 400C with a heating ramp of heating, programmed with the aim of PID that

control the energetic flux to the furnace, with a velocity of 5 degrees/min.


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MATHEMATICAL MODEL OF FIXED BEDREACTOR UNDER STEADY-STATE CONDITIONS

INTRODUCTION

In the recent years, the developing of computational power in personal computers hasextended design and performance study field, by numerical analysis, of chemical reactors.

The complexity degree of a mathematical model depending on, firstly, process type and therelative sensibility perturbation of operating conditions.Generally, for fixed bed reactors, is commonly proposed a classification as following one[9], [10] reported in Table 2

Table 2 : Classification of model for fixed bed reactors.

There are two type of mathematical model: pseudo-homogeneous model, in which is notexplicitly account the presence of catalytic bed, so hypothesizing that the differences oftemperature and concentration between solid and gaseous phase are negligible, andheterogeneous model. Inside this classification, it can be classify the mathematical model onthe base of the descriptive detail.Choice of the mathematical model that will be utilized is a delicate issue because dependupon the model complexity and upon the reached results.


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MATHEMATICAL MODEL

GEOMETRY

The reactor utilized in this work is, as described above, is a tubular fixed bed reactor withan internal diameter of 21 mm; the catalytic bed, in pellet of about 1 mm, has a lengthbetween 14 mm and 20 mm (depending on the catalyst used) and is enclosed inside twoinert regions, of the same size of catalytic bed, that has a length of about 10 cm.In Figure 11 is reported a geometric schematization of the catalytic reactor utilized.

Figure 11: Schematization of the laboratory reactor.

60 mm

25 mm

Catalytic bed ofcommercial Ni/Al 2O 3 (5g)

Inert pellets of Al2O

3

Insulating ceramicmaterial

Quartz tubularreactor

1-1.18 mm

2 mm

20 mm

Porous quartz disk

1 4 m m


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MATHEMATICAL MODEL: MASS AND ENERGY BALANCE

Mathematical model utilized for the simulation is one-dimensional, dynamic, andheterogeneous with an axial dispersion, both for heat and mass transfer, which is extendedboth in catalytic zone and surrounding inert one.In the results analysis section will be widely discussed and justified the necessity to utilize aheterogeneous description of the process (differently from often done in literature forreactor model for autothermal reforming), and will be argue the very important role, asthermal barrier, of inert packing upstream of the catalytic bed.In next paragraphs will be reported in details the model hypotheses and their analysis,

whereas in this section the mathematical equations utilized in the model are discussed.

Mass balances

During the process of autothermal reforming of methane, inside of chemical reactor, thereare the following species: CH4, O2, H2O, CO, CO2, H2 and N2. A mass balance for eachchemical specie will be write and, due to the model heterogeneity, the balance equations willbe formulated both for solid and gaseous phase.General form of mass balance, in terms of volumetric concentration, will be:

(1)

(2)

where i indicated the progressive number of the chemical specie whereas all the othersymbols are shown in the nomenclature at the end of this document.

The (1) represent the equation relative to the gaseous phase: here there is an accumulationterm, on the left to equal sign, whereas on the right, there are the axial dispersive term,

which account the flux perturbation effect induced by the presence of catalytic bed and the

)()(

,,,,int

2,

2

,,

sigigivgigi

ieff gi C C

k at C v

zC

Dt

C

=

issigigivsi r C C

k a

t

C

=

)(

)1( ,,,,


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mass diffusion in gaseous phase, the convective term and that relative to the interphasemass transfer, respectively.

The (2) instead represent the mass balance for solid phase; here there is an accumulationterm on the left to equal sign, whereas on the right, there are the interphase mass transferterm and, at last, the generation one.

The same equations are resolved also in the inert regions where, obviously, the reactionrates are zero.

Energy balance

In the following part it is reported the general form of energy balances that are utilized inthe mathematical modelling , both for solid and gaseous phase.

(3)

(4)

where j is the progressive number of investigated chemical reactions.

The eq. (3) represent the energy balance for gaseous phase; on the left to equal sign, thereis an accumulation term whereas on the right, there are the axial thermal dispersion term,due to back-mixing effect and thermal diffusion, interphase mass transfer term and, at last,the energy exchange with external environment one, respectively.

The eq. (4), instead, represent the energy balance for solid phase; on the left to equal sign,there is an accumulation term whereas on the right, there are the axial thermal dispersionterm, that contain both radiant and conductive heat transfer, the interphase energy transferterm and, at last, the generation one, respectively.

)(14

)(,,

int2

2

,

,oveng

g pgoven

r gs

g pg

f vgg

g pg

geff gT T

C h

d T T

C

ha

z

T v

z

T

C

k

t

T

+

=

)()()1( 1,,2

2

,

, j

N

j j

s p

sgss p

f vs

s p

seff s H r C T T C

ha

z

T C

k

t T R

+

=

=


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ESTABLISHMENT OF BOUNDARY CONDITIONS

The proposed mathematical model is a system of 16 differential equations, whose 9 arepartial differential equations (PDE) with variable coefficients (7 are fluid phase massbalances for the seven chemical specie and two for the energy balance of solid and gaseousphase) and 7 ordinary differential equations (ODE) (mass balances for the seven chemicalspecie in solid phase) with variable coefficients; so the problem is closed only when areestablished the relative initial and boundary conditions.

Figure 12 : Schematization of the reactor.

For mass balances the boundary conditions (for a generic chemical compound i ) are:

z=0

(5)

z=L

(6)

These conditions meaning an imposed composition at the inlet of inert region upstream ofcatalytic bed and null gradient at outlet of inert region downstream of catalyst.For the energy balance the boundary conditions are:

feed gigi C C ,, =

0,

=

z

C gi


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z=0

(7)

(8)

z=L

(9)

(10)

At the inlet of inert region upstream of catalytic bed there are a known temperature (equal

to that of alimentation) for the gaseous phase whereas there is a radiation condition for thesolid phase; in outlet of reactor instead is supposed a null temperature gradient in gas phaseand radiation condition for the solid phase.

With the regard of interface section between inert material and catalytic bed it is imposedthe continuity of the mass and enthalpy flows, for the compositions, and for thetemperatures.

ANALYSIS OF THE HYPOTHESES OF THE MODEL

In this section is reported a detailed analysis of the principal hypotheses done in the writingof the mathematical model, which is the outcome of an off-line analysis relating to:

feed gg T T =

)( 44, sgss

seff T T e zT

k =

)( 44, gsss

seff T T e zT

k =

0= z

T g


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State of gases; Analysis of the pressure drop; Kinetics in homogeneous phase; Heterogeneity;

Mass and heat dispersion in gas phase; Thermal conductibility and radiation; Determination of external heat transfer coefficient; Physical and transport properties.

STATE OF GASES

The operative conditions of pressure and temperature at which the autothermal reformingprocess of methane is conducted are:

- Pressure relatively low (experimentally lower than 2 bar):- Temperatures higher than 400 Kelvin along the length of reactor.

In these conditions, both gas and steam are considered in ideal state, so it can be write:

(11)

ANALYSIS OF THE PRESSURE DROP

With regard to the analysis of pressure drop in fixed bed reactor, it is used the Ergunequation [11], according to which the friction factor, defined from eq. (13) is equal to:

(12)

(13)

)Re1

*15075.1(*1

3 p

f

+=

2*

***

G L

d gP f pg

=

gg RT

MM P *=


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The Ergun equation for fixed bed reactors was revised from Hicks [12] whose workconcluded that Ergun equation is applicable until the following condition is satisfied:

(14)

Hicks, instead, proposed the following expression:

(15)

that is applicable for Reynolds values:

(16)

A conservative estimation of pressure drop induced from catalytic bed can be doneassuming a temperature around to 1000 K, a flow rate of 0.5 Nm^3/h, a void fraction of0.42 and a pellet diameter of about 1 mm; under these circumstances it is obtained a valueof Rep/(1- ) around to 15, so the Ergun equation is applicable; from which result an f

value around to 100 corresponding to a pressure loss of about 0.008 bar/cm, that isnegligible along the narrow length of catalytic bed.

This assumption is very common in literature and involves the absence of the equation forconservation of momentum inside the mathematical model.

KINETICS IN HOMOGENEOUS PHASE

With regard to the kinetics in homogeneous phase it is well-known that their contribute, with respect to those in heterogeneous phase, is very lower. In particular, in literature ( [13],[14] ) it is also reported an estimation, done by means of kinetic mechanism included inGRI MECH 2.11 ( [15], [16] ) calculated in conditions of partial oxidation withstoichiometric feeding.From Figure 13 , take from Bizzi et als work (2002) [13], it can be noted that the ratiobetween two reaction rates as order of magnitude is about 10^-15 in a wide range oftemperatures.

2.03

2.1

Re)1(

8.6

=

f

5000)1/(Re1000 p

500)1/(Re p


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Figure 13 : Comparison between methane conversion rates due to heterogeneous and homogeneous chem

HETEROGENEITY

With the aim to evaluate, in a preliminary phase, the limitations to thermal transportbetween solid and gaseous phase, it is applied the Mears criterion [17], that allow toestimate, comparing substantially the heat generation with that of interphase exchange one,if the temperature differences between two phase are negligible.

The relationship utilized for the application of this criterion is:

(17)

This inequality is verified where the difference of reaction rate, calculated consideringequals the temperatures of solid and gaseous bulk, are lower than 5%, condition consideredacceptable for the develop of a pseudo-homogeneous model.

a

r

f

pT

E

T d

T h

d R **15.0

*

*


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The criterion is applied in a conservative mode, considering the sole oxidation reaction, thatconstitute the conditions of highest heat generation; moreover, being the criterionapplicable to reactions with Arrhenius temperature dependent kinetics, it is estimated apseudo-activation energy for the kinetic of above mentioned reaction. The parameter valuesutilized for the application of this criterion are reported in Table 3

dp [mm] 1dr [mm] 21hf [KJ/m 2*h*K] 840PCH4 0.242PO2 0.151

rox [18], [19] T [K] 1000

Table 3: M odel parameters.

For estimation of the pseudo-activation, energy it is utilized the kinetic expression for themethane oxidation on a commercial Nickel-catalyst reported in literature, which will bediscussed in the paragraph relative to the kinetic mechanism utilized in the proposedmathematical model.

The estimation takes into account the oxidation kinetic on Nickel rather than on Rhodiumbecause it is well known that the oxidation activity on Rhodium is higher.In these conditions, Mears criterion is widely not satisfied and so it is necessary to describethe process with a heterogeneous model; according to above-mentioned, it is evident thatthis hypothesis will be more satisfied in the case of utilization of a Rhodium catalyst.For the calculation of solid-gas mass transfer coefficients value will be utilized theexpressions reported in literature [20].

(18)

for Rep < 50, or:

(19)

= 49.0Re*91.0 p D J

= 49.0Re*61.0 p D J


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for 50 < Re p < 1000

Invoking the Colburn analogy, analogous expressions are considered for the calculation ofsolid-gas heat transfer coefficients value [21].

MASS AND HEAT DISPERSION IN GAS PHASE

Turbulence mixing due to the pellets packing may be incorporated in the model consideringeffective axial dispersive coefficients in the gas phase, which include also diffusion andconduction transport phenomena, respectively for the mass and heat balance equations.

A rough estimation of the mixing effects can be done trough the calculation of the ratio

L/dp. If this ratio is higher than 50 [22], then mixing transport phenomena can beneglected. In our case, the pellet diameter is 1mm and the bed length 1cm, thus leading toL/dp equal to 10 which does not allow us to neglect mixing effects.

A more precise evaluation of the mixing relevance on the mass transport phenomena canbe done trough the calculation of the mass Peclet number, representing the ratio betweenthe rate of transport by convection and that by mass dispersion:

(20)

Considering that flow rate equal to 3 Nl/h, the real gas-phase velocity rises to values ofaround 0.7 m/s if we include the presence of the solid (e.g., with a fixed bed porosity of0.42) and the effect of high temperature in the reactor (e.g., 800 K). With a diffusioncoefficient around 3*10-4 m2/s, the mass Peclet number is slightly higher than 30. By acomparison with data reported in the literature [23] and [24], at these values of the Pecletnumber (


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(21)

valid for:

(22)

(23)

Similarly, the relevance of heat dispersion can be estimated evaluating the heat Pecletnumber:

(24)

which in our conditions is one order of magnitude lower that the mass Peclet number.

THERMAL CONDUCIBILITY AND RADIATION

For processes of partial oxidation and autothermal reforming, a very important topic forthe development of a reliable mathematical model, considering the high temperaturesand/or gradients inside of catalytic bed, is the analysis of energy transfer phenomenon.Experimental analysis and modelling of recent work in literature [26], [27], [28], [116],[117], [118] besides of the runs carried out from ourselves, show that the phenomenon ofheating of inert region upstream catalytic bed (near oxidation region) is pronounced; thisphenomenon, that will be shown in details in the following sections by IR experimentalmeasures of temperature, highlight that considering the inlet temperature at the catalytic

])

Re7.9

1(

5.0Re

*73.0[,

i p

i p piieff

ScSc

d v D

++

=

50Re008.0 p

mmd p 6377.0

in t ,

,

* * *

*g p g

he ff g

v L C P e

k M M

=


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bed equal to that of reagent feeding is wrong because this value, is underestimated, in somecases, also of hundreds degrees.In fact, from these considerations, is born the necessity to develop a mathematical modelthat is expanded to inert regions in contact with the catalytic bed.

The evaluation of this phenomenon is included, mathematically, inside the axial thermaldispersion coefficient.In particular, for the gas phase, the thermal phenomena that have to be consider are theconductive and radiant ones and that due to backmixing; whereas for the solid phase, infixed bed, have to be consider the radiant and conductive thermal ones.Obviously, the experimental evaluation of these phenomena oddly is rather difficult; at thispurpose, there are many theoretical and experimental works for evaluation of the effect ofaxial thermal dispersion in fixed bed [29], [30], [31], [32] S. Yagi,and D. Kuni, Aiche J. 6,543, (1960).[33], [34], [35] R. Krupiczka, Chim. Ind. Genie Chim.; 95 (6), 1393, (1966).[36], [37], [38], [39], [40].In particular, the first values of axial thermal conductibility in fixed bed were obtained by

Yagi et al [32] by means of experimental measures of axial temperature along a fixed bedheated by an infrared lamp and crossed from a known countercurrent air stream; theinterpolation of these measures with a conductive-convective mathematical model led tothe determination of parameter of interest.

The experiments, carried out on different materials and dimensions of the catalytic bed ledto determination of following expression:

(25)

This expression is then verified from ulterior experimental analysis of other Authors, also with different materials and shapes of the constitutes of the bed, determining its validity fora Reynolds number higher than 0.8.

The first term of this expression, also shown from previous work [29], [30], [34], representthe effective axial thermal conductibility value for beds in a stagnant flow, including alsoconductive phenomena that interesting substantially the solid phase; commonly it is utilizedKrupiczkas expression [35] R. Krupiczka, Chim. Ind. Genie Chim.; 95 (6), 1393, (1966).[36] according to which:

(26)

RePr *0, += e f axeff k k k

n

f

s

f

e

k k

k k

=

0


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where:

(27)

The experiments are carried out with low temperatures and in absence of high temperaturegradients so, it is evident, that in the above expression are considered conductive,convective and phenomena but not the radiation ones.Instead, when the bed is submitted to high thermal levels and gradients, as well in

autothermal process for the hydrogen production, it is necessary consider also the radianteffects. With this purpose, Wakao and Kato [37] proposed the following expression:

(28)

where

(29)

valid for

(30)

(31)

( )

=

f

s

k k n log057.0log757.028.0

100020 f

s

k k

3.0r Nu

)() ( 000 RADIANT k CONDUCTIVE k k ee e +=

11.1

96.00 707. 0)(

= sr e

k k

Nuk RADIANT k


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and where

(32)

and

(33)

For the presence of water and carbon dioxide, also for the gas phase would haveconsidered the radiant phenomenon; with this purpose Wakao [41] proposed also amodified expression of h r for gas phase. But it has to be considered that the emissivity ofgaseous compounds is strongly dependent from temperature, void dimension and partialpressure, that in fixed bed generally have very low value and so, in this work, it is neglected.For example, the emissivity of CO 2 at high temperature (about 1200 K) and in a void radius

of 1 cm has a value of just 0.05 [42].

DETERMINATION OF THE EXTERNAL HEAT TRANSFER COEFFICIENT

In the proposed mathematical model there is, in the equation ( 3) , the external heat transferterm and it is necessary to know global external heat transfer coefficient h OVEN . Thedetermination of this value is done off-line by means of experimental measures opportunelyrealized. The energy balance, in steady-state, for a non-reactive fixed bed reactor is writtenas following:

(34)

s

pr r k

d h Nu =

3

100264.02

2268.0

= T

e

hr

)(14

0,

int2

2

,

,oveng

g pgoven

r

gg

g pg

geff T T C

hd z

T v

z

T

C

k

=


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So, if the reactor is submitted to a flow of non-reactive mixture (with known physical andtransport properties) and measuring the thermal profile inside the bed, it is possible, if it isknown the external temperature, determine the global heat transfer coefficient.On the basis of these observations, experimental runs that utilized the same reactor ofreforming runs (with the same insulating ceramic material) pose inside the furnace with aprefixed temperature and submitted to a known nitrogen flow (non-reactive mixture). Thereactor was equipped with a mobile thermocouple, by means a quartz coaxial sleeve, for themeasure of axial thermal profile (in this type of experimental run it is possible utilize athermocouple for measure of thermal profile because in this non-reactive system thetemperature differences between the two phases are equal to zero).

With the above discussed experimental protocol, two experimental runs has been realized with different temperature of the external oven (623 e 773 K respectively) measuring, at thesteady state conditions, the thermal profile.

In Figure 14 , the read and blue points indicate the experimental results so obtained. Then they are fitted, based on equation ( 34) utilizing as unknown parameter the globalexternal heat transfer coefficient. From the Figure 14 (dotted and continuous lines) it canbe noted that the results of both the experimental runs are satisfactorily fitted from theproposed equation with a value of h oven equal to 9 W/m^2*K. It is necessary highlightedthat the global external heat transfer coefficient results independent from the temperatureused in the studied range.

Axial position [cm]

0 2 4 6 8 10 12 14 16

T e m p e r a

t u r e

[ K ]

400

450

500

550

600

650

700

750ExperimentalSimulatedSimulatedExperimental


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Figure 14 : Axial thermal profile at steady state conditions: pressure:1.7 bar;; nitrogen flow: 4 Nl/min; oven tem623 and 773 K.

The experimental value of global heat transfer coefficient will be utilized, in this work, forall the results that will be obtained.

PHYSICAL AND TRANSPORT PROPERTIES

Inside proposed mathematical model have to be defined the kinetic law and all the physicaland transport properties of gaseous and solid phases.

The kinetic law, that is strongly dependent of nature of the catalyst, will be discussed in thefollowing sections that will reported the results obtained with both Nickel and Rhodiumcatalytic systems, whereas the physical and transport parameters will not be here written butare all function of temperature and composition, and so they are variable along all thereactor, and are calculated by means of well-known relationship commonly utilized inchemical engineering field [20].

NUMERICAL METHOD

The mathematical model constituted from the differential equations of materials and heat

balances (16 differential equations of which 9 with partial derivative PDE and 7 ordinarydifferential equations ODE all with variable coefficients), have to be numerically solvedso it is necessary approximate the problem with differential formula [60].From the solution of these approximated equations it can be obtained scalar values ofunknown functions, that is a series of values that correspond to a set of points on thedomain. These values are the scalar unknown quantity of the approximate problem with

which the real problem has been substituted. The differential equations with partial derivative represent a good relation on all the pointsof the integration domain. For this reason, it is possible write, for each point an equation aslong as the partial derivatives are expressed in function of the scalar unknown quantity.

The expressions of the partial derivative approximate from functions with scalar unknownquantity determine the solve method adopted, implicit or explicit. Moreover, dependingfrom how the partial derivative are expressed, there are discretization errors with respect tointegration step in space and in time.

The truncation error that represent the difference between the solution of the startingdifferential equation and its approximation, depend on also form the truncation error(round off) related with finite dimension of the machine registry.


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Another important information is that about the definition of local and global truncationerror: the first correspond to the difference between the exact solution starting from theprevious step and the calculated value, whereas the second is the difference between theexact solution and the calculated value. In conclusion, the global error is due tocombination of the different local errors but is not really the arithmetic sum.

The amplification of the errors in the solutions represents the so-called instabilityphenomenon of numerical method chosen. A method is defined stable if the differencebetween the exact solution of the initial problem (without approximation) and numericallycalculated value (with approximation) not diverge for infinite time. The stability of anumerical method depend on both the solve method and from the form of startingdifferential equation, so fixed the sample differential equation that have to be solved, it ispossible evaluate the extreme stability of a method.

DISCRETIZATION OF THE SYSTEM

From previous discussions it is gathered that to obtain a model that is numerically solvedthe system have to be divided in nodes, that are a series of volumes with little but finitedimensions [61].

The nodes are numerated, in the space, along the axis of the system ( Figure 15 ) in whichthe inlet and outlet nodes (respectively 0 and n+1) are simply nodes of convectivetransport.

Figure 15 : Schematization of the spatial discretization of the system.

After the discretization of the system in the space, is possible formulate for each node adifferential equation that will be of ordinary type because it is differentiable only in the time.

1 2 N-1 N N+1

CATALIZZATORE


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The spatial discretization method chosen is backward finite difference formula, that is innode i the expression of first derivative (in this case there are not diffusive terms and sothere are not second derivatives) will be approximated:

)( 21 zO z

y y

z y ii

i

+=

(35)

where y represent the generic unknown function and i the nodes studied.So the model is constituted of a set of ordinary differential equations (ODE) that are solvedby means of numerical methods with a computer.

With this aim it was chosen a solver found in the library of the software ofWolframResearch Mathematica that is named NDsolve.

The syntax of this solver is the following:

NDsolve[{y 1 [t]==y 1 [t],,y n [t]==y n [t]},{y 1 [0]==y 0

1 ,,y n [0]==y 0

n },{y 1 [t],,y n [t]},{t,t min ,t max }]

In which, it is assigned the vector of the differential equation system, the vector of the ofthe starting conditions, the solution functions one and finally the integration time span;then the software calculate the numerical solution of even system utilizing as integrationmethod the explicit Runge-Kutta method which order is automatically managed from thesolver (this is a default option but it is possible a manually management).


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RESULTS ON NICKEL BASIS CATALYST

The principal aim of the present work is to validate the proposed mathematical model bycomparing the calculated results with the experimental data obtained during the runsperformed with the experimental device previously presented.In particular, in the present chapter, the data obtained using a commercial Nickel catalyst

will be shown. The performed study can be summarized in the following steps:

Thermodynamic analysis for the evaluation of the experimental conditions that not allo formation;

Analysis of the kinetic mechanism of the autothermal reforming of methane on Nickel cat Numerical-experimental prediction of the overlapping between the oxidative and the refor Runs at different flow rates and compositions Numerical simulations and comparison with the experimental data Numerical- experimental considerations on transport phenomena and model limitations Conclusions for CPO and ATR processes on commercial Nickel catalyst.

Each reported simulation it is performed at a starting bed temperature higher then 750 K(near the activation temperature of the catalyst), as previously said in the experimentalsection, and besides it is hypothesized a start-up with an empty reactor, in other word it ishypothesized a vessel initially pressurized with nitrogen (no reactant at initial conditions).

This situation does not influence the process kinetic so the vessel, at the start-up, can beconsidered in stationary conditions.

THERMODYNAMIC ANALYSIS

The experimental program is performed making a thermodynamic preliminary study toobtain the feed ratio (Air/CH4) that not allows the coke production. In fact the cokeformation is undesired and, besides, is not foreseen by the kinetic models adopted.It is important to note that the thermodynamic information reported in the presentparagraph, independently from the catalyst are considered a pattern also for the compilation


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of the experimental program on the commercial Rhodium catalyst shown in the followingsection.

The thermodynamic analysis of the target process is realized using Aspen Plus, the softwareof the Aspen Technology [ 65 ] .In particular, the simulations are conducted by adopting the following characteristicsincluding in that software:

Reactor RGibbs type; Property method Peng-Robinson.

In details, the RGibbsreactor allows the determination of the thermodynamic equilibriumconditions by means the minimization of the Gibbs free energy; this model can be used for

the following equilibriums:

Single phase (liquid o steam); Steam phase in presence of different liquid phases; Solid phase in solution; Conventional solid compounds in liquid phase.

The property method Peng-Robinson is commonly used in the thermodynamic analysis of polarand non-polar mixtures, in particular for light hydrocarbons and gases.From a literature survey [1] it is well known that the addition of water to the feed of partialoxidation processes, from a thermodynamic point of view, limits the channel blackproduction. So it is clear that a conservative analysis direct to the determination of the feedratio A/CH4 that allows this conditions have to be realized in the case of partial oxidation.

The operative conditions adopted in the simulation are summarized in Table 4

Reactor type RGibbsExternal duty nulloPreheating temperature [K] 623Pressure [bar] 2Feed ratio Air/CH4 2-5Feed ratio H2O/CH4 0

Table 4: Range of the parameters used for the thermodynamic analysis.


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In Figure 16 the results of the simulations obtained with Aspen Plus are shown.In particular, the yield in CO, H2 and C (graphitic) are reported, the expressions of theseparameters are written below:

(36)

(37)

(38)

where y i are the equilibrium compositions.From that figure it is clear that the thermodynamics foresees the absence of coke for valuesof A/CH4 higher then 2.8 and a non monotonous evolution of yield in syngas with amaximum in correspondence of value a little bit higher than 3.

A/CH4

2.0 2.5 3.0 3.5 4.0 4.5 5.0

R e s a

C , R

e s a

H 2

, R e s a

C O

0

20

40

60

80

100

A/C vs resaC A/C vs resaH2 A/C vs resaCO

42 CH COCOC

C

y y y y y

YieldC +++

=

42 CH COCOC

CO

y y y y y

YieldCO+++

=

42 CH COCOC

CO

y y y y y

YieldCO+++

=

8/9/2019 Tesi Dottorato S

Tesi Dottorato Scognamiglio

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